Catalytic sulfur removal from a hydrocarbon stream

ABSTRACT

A process for the catalytic removal of sulfur from a hydrocarbon stream such as gasoline comprising an organo-sulfur compound such as a mercaptan or thiol. The catalyst is a silica based zeolite such as ZSM-5. The process is preferably performed in a downer reactor with a residence time of between 7 and 30 seconds and a volumetric particle concentration of between 15 and 40%. Preferably, substantially all of the sulfur that is removed from the organo-sulfur compound is in the form of hydrogen sulfide.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims the benefit of U.S. provisional patent application Ser. No. 60/632,560 filed Dec. 2, 2004, the entire contents of which are hereby incorporated by reference.

FIELD OF THE INVENTION

The invention relates to the catalytic removal of sulfur from hydrocarbon streams comprising a sulfur containing hydrocarbon. In particular, the invention relates to the removal of sulfur compounds such as mercaptans and thiols from hydrocarbon streams such as gasoline by catalytic conversion to hydrogen sulfide using a zeolite catalyst such as ZSM-5. The invention may be practiced using a solids transport reactor and is advantageously practiced in a downer reactor.

BACKGROUND OF THE INVENTION

Sulfur in fuels for internal combustion engines is generally undesirable. In the auto industry, high sulfur levels in gasoline lead to poisoning of catalytic converters and to corrosion caused by sulfuric acid compounds. Sulfur in fuel is also detrimental to the environment. Sulfur dioxide (SO₂) emissions can be traced directly to the combustion of sulfur species in transportation fuels. SO₂ emissions are of particular concern given that they are precursors of acid rain and sulfate aerosol formation, which contribute considerably to total ambient fine particulate matter. Recent environmental regulations establish strict sulfur levels in transportation fuels. As a result, the refining industries are under constant pressure to achieve more rigorous standards of product specification for sulfur.

Fluid Catalytic Cracking (FCC) is used to produce gasoline. FCC gasoline contributes from 25% to 40% of the total volume of the “gasoline pool”. Gasoline coming from the FCC unit may have sulfur levels as high as 2000 ppm, and FCC gasoline accounts for more than 90% of the total sulfur in the “gasoline pool”. Refineries must therefore focus on reducing sulfur in FCC gasoline to achieve the target sulfur levels.

The hydrotreating of FCC gasoline, or hydrodesulfurization (HDS), is a post-FCC treatment process performed to achieve low sulfur levels in gasoline. Catalysts used are CoMo/γAl₂O₃ and NiMo/γAl₂O₃. Hydrodesulfurization reactions are exothermic and occur simultaneously with hydrogenation reactions. Hydrogenation reactions decrease the quality of gasoline since the olefins, which contribute to the gasoline octane number, are converted into alkanes (paraffins). It is well known that there are no conditions in the HDS process where hydrogenation can totally be excluded. During HDS, sulfur is removed in the form of hydrogen sulfide (H₂S); however, H₂S reacts with olefins in the fuel, producing mercaptans. The use of hydrogen adds a significant cost to gasoline production. Thus, a process that does not require hydrogen to remove sulfur from FCC gasoline is desirable.

Some attempts have been made to address the need for an improved process for removing sulfur from fuels. Methyl mercaptan (methane thiol) conversion was studied in a fixed bed reactor containing H-ZSM5 at 482° C. and 1 h⁻¹ LHSV obtaining conversions of 99% (Chang C. D., Silvestri A. J., J. of Catalysis, Vol. 47, 1977, pp 249-259). The products obtained were hydrocarbons ranging from C₁ to C₁₁ and H₂S. About 27% of the carbon feed was undesirably converted to dimethyl sulfide, which has no commercial value in fuels. This process is not commercially viable due to the long residence time and the detrimentally high rate of conversion of gasoline to dimethyl sulfide.

Collins et al. in U.S. Pat. No. 5,401,391 and U.S. Pat. No. 5,482,617 disclose a dense phase fluidized bed reactor with H-ZSM5 catalyst used to desulfurize hydrocarbons. Organic sulfur compounds were converted to H₂S. The catalyst used contained 25% of H-ZSM5 zeolite with a Si/Al molar ratio of 25. The process could be used to desulfurize either light gases or gasoline streams with light olefins being upgraded to more valuable gasoline range materials. Reaction conditions were: 370-450° C., 50-250 psig, and 0.1-2 h⁻¹ WHSV. Up to 61% of sulfur in the feed was converted to H₂S. This level of conversion is not high enough to be commercially useful as an alternative to hydrotreating. The low overall conversion may be due in part to the significant non-uniformity of fluid distribution of the gas phase in a dense phase fluidized bed. Dense phase fluidized beds normally have catalyst residence times on the order of minutes and gas phase residence times on the order of a few seconds. The gas phase residence time in a dense phase fluidized bed is generally too short to achieve significant levels of conversion with minimal amounts of intermediate organo-sulfur compounds, especially when considering a reaction network involving a series-parallel reaction mechanism such as in the present invention. In addition, the long residence time of the catalyst in the dense phase fluidized bed leads to significant irreversible adsorption of reactants and reaction products on the catalyst and causes the undesirable formation of aromatic hydrocarbons on the catalyst. Aromatic hydrocarbons negatively impact gasoline quality and are pre-cursors to coke formation. Some aromatics, such as benzene, are undesirable carcinogenic species. Aromatics have to be capped in gasoline at levels of less than 10%. As a result aromatics are preferably avoided in commercial gasoline production.

Desulfurization of thiophene on H-ZSM5 was studied in a mini-fixed bed reactor using alkanes as co-reactants at 400-500° C. (Yu, S. Y.; Waky, T.,; Iglesia, E., Appl. Catal. A, Vol. 242, 2003, pp. 111-121). Propane, n-hexane, and n-decane were used as co-reactants. It was found that thiophene desulfurization increased with increasing alkane chain size. It was observed that thiophene does not alter the nature of alkane reaction pathways on H-ZSM5, but increases the selectivity to aromatics as a result of the selective reaction of alkane-derived species with unsaturated fragments formed during thiophene decomposition and thiophene desulfurization. The aforementioned problems of fluid distribution, mass transfer limitations, low overall conversion, and adsorption of reaction products leading to reduced gasoline quality and coke formation are all disadvantages of fixed bed systems when used for catalytic de-sulfurization.

The need therefore still exists for an improved process for catalytic removal of sulfur from hydrocarbon streams.

SUMMARY OF THE INVENTION

According to one aspect of the invention, there is provided a process for removing sulfur from a hydrocarbon stream comprising an organo-sulfur compound, the process comprising: providing a downer reactor; providing a hydrocarbon stream comprising an organo-sulfur compound in the downer reactor at a temperature in the range of 300 to 500° C.; providing a catalyst comprising a zeolite that promotes removal of sulfur from the organo-sulfur compound in the downer reactor; contacting the hydrocarbon stream with the catalyst; and, converting at least 70% of the organo-sulfur compound to hydrogen sulfide in the downer reactor.

According to another aspect of the invention, there is provided a process for removing sulfur from a hydrocarbon stream comprising an organo-sulfur compound, the process comprising: providing a solids transport reactor; providing a hydrocarbon stream comprising an organo-sulfur compound in the solids transport reactor at a temperature in the range of 300 to 500° C.; providing a catalyst comprising a zeolite that promotes removal of sulfur from the organo-sulfur compound in the solids transport reactor; the catalyst having a catalyst reactor density from 0.045 to 0.12 g_(cryst)·s·cm⁻³; contacting the hydrocarbon stream with the catalyst; and, converting at least 70% of the organo-sulfur compound to hydrogen sulfide in the solids transport reactor in a residence time from 7 to 20 seconds, wherein the product of catalyst reactor density and residence time is from 0.36 to 1.44 g_(cryst)·s·cm⁻³.

According to yet another aspect of the invention, there is provided a process for removing sulfur from a hydrocarbon stream comprising an organo-sulfur compound, the process comprising: providing a downer reactor having a length of from 10 to 36 meters; providing a hydrocarbon stream comprising gasoline and comprising 10 to 10000 ppm of an organo-sulfur compound comprising a mercaptan in the downer reactor at a temperature in the range of 300 to 500° C.; providing a catalyst comprising zeolite ZSM-5 in the downer reactor at a volumetric particle concentration of from 5% to 40%; contacting the hydrocarbon stream with the catalyst; and, converting at least 90% of the organo-sulfur compound to hydrogen sulfide in the downer reactor in a residence time of between 7 and 20 seconds.

Surprisingly, the process of the present invention provides high conversion of sulfur compounds to hydrogen sulfide at relatively short residence times with no undesirable thio-paraffin intermediates such as diethyl sulfide. Short residence times decrease the size of the reactor, improving the economics of the process, and reduce the likelihood for secondary olefin condensation reactions on the catalyst. This in turn reduces the production of aromatic hydrocarbons, which desirably limits any detrimental effect of the process on gasoline quality and limits the formation of coke on the catalyst. The process of the present invention is particularly advantageously carried out in a downer reactor, which permits optimization of process conditions over a range that achieves desirably high conversion while maintaining high selectivity.

The process of the present invention may be operated in a regime that provides the following key advantages:

-   1. High conversion of orgno-sulfur compounds using only a catalyst,     obviating the need for hydrogen; -   2. High selectivity of converted organo-sulfur compounds to H₂S; -   3. High levels of olefins due to negligible gasoline cracking and     the surprising formation of low molecular weight olefins due to     intra- and inter-molecular dehydrosulfidation; -   4. Negligible olefin condensation and coke formation; and, -   5. All of the above achieved within practical reactor sizes and     process conditions.

The hydrocarbon stream may comprise from 1 to 20 carbon atoms and may comprise alkanes, alkenes, alkynes, aromatics, or a combination thereof. Preferably, the hydrocarbon stream comprises a mixture of paraffins, olefins, cyclo-paraffins and aromatics having from 6 to 12 carbon atoms. The hydrocarbon stream may comprise, for example, n-octane. The hydrocarbon stream may comprise gasoline or light cycle oils such as diesel fuel. Most preferably, the hydrocarbon stream is gasoline. The gasoline may be the product of an FCC process.

The organo-sulfur compound may comprise one or more sulfur atoms and from 1 to 12 carbon atoms. Preferably, the organo-sulfur compound comprises one or two carbon atoms and from 2 to 10 carbons atoms, more preferably from 2 to 8 carbon atoms, more preferably from 2 to 4 carbon atoms. The sulfur may be bonded to one or two carbon atoms. The organo-sulfur compound may comprise a mercaptan, a thio-ether, a thio-ketone, a thio-aromatic, a thio-paraffin, or a combination thereof. Preferably, the organo-sulfur compound comprises methyl mercaptan, ethyl mercaptan, or diethyl sulfide. The organo-sulfur compound may be present in the hydrocarbon stream at a concentration of from 10 to 10000 ppm, preferably from 50 to 5000 ppm, more preferably from 75 to 2000 ppm, still more preferably from 100 to 1000 ppm, yet more preferably from 120 to 500 ppm, even more preferably from 150 to 350 ppm.

The temperature in the reactor is any suitable temperature for promoting catalytic dehydrosulfidation of the organo-sulfur compound. Dehydrosulfidation is the process by which sulfur is removed from an organo-sulfur compound as hydrogen sulfide, for example, in a manner similar to the removal of oxygen from methanol as water in a de-hydration process. Generally, higher temperatures lead to faster reaction rates, but favour undesirable catalytic cracking of the hydrocarbon stream. There is therefore an optimal temperature range for the dehydrosulfidation process. The temperature may be from 300 to 500° C., preferably from 350 to 475° C., more preferably from 375 to 470° C., still more preferably from 400 to 460° C., even more preferably from 425 to 455° C., most preferably about 450° C.

The catalyst comprises a zeolite that promotes removal of sulfur from the organo-sulfur compound, preferably through dehydrosulfidation. The catalyst includes the zeolite crystallites and may include a supporting matrix. Any suitable zeolite crystallite or combination of crystallites may be used as catalyst. The catalyst may include noble metals, transition metals or transition metal oxides. The catalyst may include phosphorous. The catalyst is preferably a microporous hydrophobic zeolite having acidic sites and a tortuous pore path with a pore diameter of from 5-7 Å, preferably about 5.4 Å. The catalyst may comprise a silico-aluminate comprising silica in an amount of greater than 85% by weight, preferably greater than 90% by weight, more preferably greater than 95% by weight. Most preferably, the catalyst comprises the zeolite ZSM5, which is understood to include H-ZSM5. The catalyst may comprise a combination of different zeolites.

The hydrocarbon stream is contacted with the catalyst in the reactor. Preferably the reactor is a solids transport reactor, for example a circulating fluidized bed, a riser, or a downer. Most preferably, the reactor is a downer. The length of the downer may be from 5 to 40 m, preferably from 10 to 36 m, more preferably from 12 to 30 m, still more preferably from 14 to 27 m. The fluid velocity of the hydrocarbon stream in the reactor may be from 0.5 to 4.0 m/s, preferably from 1.0 to 3.0 m/s, more preferably from 1.5 to 3.0 m/s, still more preferably from 1.8 to 2.8 m/s.

The catalyst circulates through the reactor with the fluid and has a residence time within the reactor. The catalyst residence time and the fluid residence time in the reactor are similar, preferably within about 20% of one another. The catalyst residence time should be less than about 30 seconds to prevent undesirable coke formation on the catalyst and should be greater than about 7 seconds to provide the desired level of conversion while allowing reasonable quantities of catalyst to be used. The catalyst residence time may be from 7 to 30 s, preferably from 8 to 20 s, more preferably from 9 to 18 s, still more preferably from 12 to 15 s.

Generally, conversion is related to residence time and catalyst reactor density, which is the mass of active catalyst crystallites in a given volume of reactor. The catalyst reactor density is preferably from 0.02 to 0.12 g_(cryst)·cm⁻³, more preferably from about 0.04 to 0.10 g_(cryst)·cm⁻³, still more preferably from about 0.05 to 0.09 g_(cryst)·cm⁻³ The product of residence time and catalyst reactor density is a constant for a given level of conversion. Preferably, the product of residence time and catalyst reactor density is from 0.30 to 1.44 g_(cryst)·s·cm⁻³, more preferably from 0.60 to 1.00 g_(cryst)·s·cm⁻³, still more preferably from 0.65 to 0.85 g_(cryst)·s·cm⁻³, yet more preferably from 0.70 to 0.80 g_(cryst)·s·cm⁻³.

The level of conversion of the organo-sulfur compound in the reactor is at least 70%, preferably at least 75%, more preferably at least 80%, still more preferably at least 85%, yet more preferably at least 90%, even more preferably at least 95%, still even more preferably at least 97%, most preferably at least 99%. Catalyst reactor density and residence time may be adjusted within the ranges provided above to achieve the desired level of conversion in the reactor. Preferably, substantially all of the sulfur that is removed from the organo-sulfur compound is in the form of hydrogen sulfide. However, not all of the sulfur in the organo-sulfur compound is necessarily converted to hydrogen sulfide; some of the sulfur may remain in the organo-sulfur compound (for example, thio-aromatics) or alternative reaction products (for example, sulfur containing coke) may be formed in the dehydrosulfidation reaction.

The volumetric concentration of particles, which is the percentage of the reactor volume occupied by catalyst particles (matrix and crystallites), may be selected to achieve the desired conversion at a given residence time. The volumetric particle concentration is related to the catalyst reactor density by the apparent density of the catalyst particles. The volumetric concentration of catalyst should not be so high as to create non-uniformity of fluid flow, particularly choking in a riser reactor. The volumetric particle concentration may be from 5 to 40%, preferably from 8 to 30%, more preferably from 10 to 25%, still more preferably from 14 to 23%.

When the hydrocarbon stream contains a mixture of paraffins, olefins, and aromatics, and particularly when the organo-sulfur compound is a thio-aromatic, there may be a tendency for coke formation on the catalyst. The catalyst may be regenerated through combustion of coke. The regenerator may be a dense phase fluidized bed reactor. The catalyst may circulate continuously between the solids transport reactor and the regenerator.

Further features of the invention will be described or will become apparent in the course of the following detailed description.

BRIEF DESCRIPTION OF THE DRAWINGS

In order that the invention may be more clearly understood, embodiments thereof will now be described in detail by way of example, with reference to the accompanying drawings, in which:

FIG. 1 shows conversions of ethyl mercaptan versus time with a feed composition of 10 wt. % (10000 ppm) EM in nC₈;

FIG. 2 shows conversions of ethyl mercaptan versus time with a feed composition of 5 wt. % (5000 ppm) EM in nC₈;

FIG. 3 shows conversions of n-octane versus time with a feed composition of 100 wt. % nC₈;

FIG. 4 shows conversions of n-octane versus time with a feed composition of 10 wt. % EM in nC₈;

FIG. 5 shows conversions of n-octane versus time with a feed composition of 5 wt. % EM in nC₈;

FIG. 6 shows the relationship between catalyst density and reaction time for a conversion of 95% at 450° C.; and,

FIG. 7 shows a schematic representation of an embodiment of a process according to the present invention.

DESCRIPTION OF PREFERRED EMBODIMENTS

Ethyl mercaptan reacts over H-ZSM5 as follows:

-   1. Intra-molecular dehydrosulfidation. Ethyl mercaptan reacts via     intra-molecular dehydrosulfidation to give ethylene and H₂S,     CH₃−CH₂SH⇄CH₂=CH₂+H₂S  (1) -   2. Inter-molecular dehydrosulfidation between two mercaptan reacting     molecules. This reaction leads to the removal of H₂S molecule from     two mercaptan molecules yielding diethyl sulfide (DiE-S) and H₂S:     CH₃−CH₂SH+CH₂SH−CH₃⇄CH₃−CH₂−S−CH₂−CH₃+H₂S  (2)     -   Moreover, following, this first step, a second step involves         further intra-molecular dehydrosulfidation of diethyl sulfide         (DiE-S) yielding an olefin (butene) and H₂S:         CH₃−CH₂−S−CH₂−CH₃⇄CH₃−CH₂=CH₂−CH₃+H₂S  (3)

Equilibrium constants and equilibrium compositions of the proposed set of three simultaneous reactions for the dehydrosulfidation of ethyl mercaptan (EM) can be considered at set operation conditions: temperature, pressure, reactant concentration. The theoretically calculated chemical equilibrium constants for the three reactions are reported in Table 1. Reactions 1 and 3 are endothermic reactions while reaction 2 is exothermic. All equilibrium constants are higher than one, thus all three reactions are favored under the selected operating conditions.

Regarding the three reactions involved it should be mentioned that reactions 1 and 2 are of the competitive type while reactions 2 and 3 are in series. Thus, the overall reaction scheme of dehydrosulfidation is a combined parallel-in series reaction called, in the context of this invention, “parallel-series” reaction network with the three main reaction steps contributing to the final product distribution. TABLE 1 Equilibrium constants in the 350-500° C. temperature range for the EM dehydrosulfidation reactions (equations 1, 2 and 3). T [° C.] K₁ (eq. 1) K₂ (eq. 2) K₃ (eq. 3) 350 4.4 3.0 412.7 375 8.1 2.9 695.0 400 14.7 2.8 1136.2 425 25.4 2.8 1808.8 450 42.8 2.7 2811.9 475 70.0 2.7 4278.6 500 111.7 2.6 6385.6

Observing the calculated equilibrium constants it can be noticed that for the temperature range of interest 350-500° C., the equilibrium constants for reaction 1 are consistently higher than the equilibrium constants for reaction 2, and the equilibrium constants for reaction 3 much higher than those for reactions 1 and 2. This is an indication that diethyl sulfide species are formed and consumed readily in all cases.

Calculations of expected conversion indicate that high conversion of mercaptans coupled with elimination of the high levels of di-sulfide observed in prior art processes should be achievable. However, the exact conditions needed to achieve the desired results is not evident. In order to design a process that achieves the desired conversion, an overall kinetic rate constant should be determined experimentally under conditions similar to those expected in actual production. The distribution of reaction products should also be examined to determine whether the gasoline is being cracked in the process of removing sulfur from the hydrocarbon stream. The undesirable formation and accumulation of aromatics should also be monitored. The goal is to determine a set of process conditions that achieves high conversion of organo-sulfur compounds to H₂S without cracking gasoline that can be implemented in a practically sized reactor.

Experimental Protocol

Experiments were conducted using the CREC Catalytic Simulator described in U.S. Pat. No. 5,102,628 (de Lasa), which is hereby incorporated by reference. The 52 cm³ CREC Catalytic Simulator consists of two outer shells, lower section and upper section that permits to load or to unload the catalyst easily. This reactor was designed in such way that an annular space is created between the outer portion of the basket and the inner part of the reactor shell. A metallic gasket seals the two chambers, an impeller located in the upper section. A packing gland assembly and a cooling jacket surrounds the shaft supports the impeller. Upon rotation of the shaft, gas is forced outward from the center of the impeller towards the walls. This creates a lower pressure in the center region of the impeller thus, inducing flow of gas upward through the catalyst chamber from the bottom of the reactor annular region where the pressure is slightly higher. The impeller provides a fluidized bed of catalyst particles as well as intense gas mixing inside the reactor.

The CREC Catalytic Simulator operates in conjunction with a series of sampling valves that allow, following a predetermined sequence, to inject hydrocarbons and withdraw products in short periods of time. An Agilent 6890N gas chromatograph (GC) with flame ionization detector (FID) and a mass selective detector (MSD, Agilent 5973N) allows the quantification of reaction products using a capillary column HP-5 Phenyl Methyl Siloxane with a length of 30 m, a nominal diameter of 0.32 mm, and a nominal film thickness of 0.25 μm. More details of the mechanical design of the Catalytic Simulator are given by Kraemer (Kraemer, D., Modelling Catalytic Cracking in a Novel Riser Simulator, Ph.D. Thesis, University of Western Ontario, London, Ontario, 1991) and Pruski (Pruski, J., Adsorption Phenomena During FCC in a Novel Riser Simulator, M.E.Sc. Thesis, University of Western Ontario, London, Ontario, 1996).

A catalyst comprising ZSM-5 crystallites prepared following the method of Gabelica, et al. (Gabelica, Z.; Blom, N.; Derouane, E. G. “Synthesis and Characterization of ZSM-5 Type Zeolites”, Appl. Catal, Vol. 5, 1983, pp. 227-248) was provided in the CREC Catalytic Simulator. No binder was utilized with the catalyst. The total mass of catalyst crystallites was about 0.2 g.

Ethyl mercaptan C₂H₆S (Alfa Aesar, CAS number 75-08-1), was selected as a key chemical species to evaluate the mercaptan conversion and to assess the reaction network. A hydrocarbon stream comprising gasoline was simulated with n-octane (EM Science, CAS number 111-65-9), a straight-chain hydrocarbon with a boiling point falling in the middle of the gasoline boiling range.

Mixtures of these compounds were reacted at different concentrations (5 and 10 wt. %), at three temperatures (350, 400 and 450° C.) and at four different contact times (10, 20, 40 and 60 s), using a catalyst to oil ratio of C/O=2.5. The total mass of the injected stream was 0.08 g. All experiments were repeated at least 3 times to secure reproducibility of results. These conditions were selected as considered representative of the ones that could be encountered in a potential industrial post-treatment process of gasoline dehydrosulfidation.

Three type of experiments were conducted:

-   1. Thermal runs with pure n-octane, -   2. Thermal cracking runs with 90% n-octane and 10% ethylmercaptan, -   3. Catalytic runs with ZSM5 and 90-95% n-octane and 5-10%     ehtylmercaptan     Thermal Runs Using Pure n-Octane

To assess the possible thermal effects on n-octane, the paraffin hydrocarbon species used to model gasoline, thermal runs were developed in the CREC Catalytic Simulator using a reaction time of 60 s and three temperatures: 350° C., 400° C. and 450° C. A sample of 0.08 g of pure n-octane was injected into the reactor in each run. These conditions were studied to evaluate the influence of thermal reactions under severe process conditions (e.g. highest possible conversions).

Results of the thermal experiments are reported in Table 2. For each condition, both conversion and mass balance closure are average values for three repeat runs. Experiments developed with pure nC₈ at both 350° C. and 400° C. and 60 s, the largest reaction time studied, showed no chemical species out of nC₈. This result demonstrates that at 400° C. and lower thermal levels, there is no significant nC₈ thermal conversion. Furthermore, at 450° C. and 60 s, there is an indication of a very small amount of nC₈ being converted to ethane and propene, with this fraction being limited to 0.34%. In this product fraction ethene is present in a larger fraction than propene (0.21 and 0.13 wt. %, respectively, which is about a factor of two) with this product distribution being characteristic of thermal cracking where dominant β-scission cracking promotes ethene as the more abundant product. Table 2 also reports mass balance closures with these balances that are well in the range of typical closures achieved in the CREC Catalytic Simulator.

Thus, on the basis of the data reported, thermal cracking of gasoline model species is neglected at 450° C. and 60 s. TABLE 2 Conversion of n-octane at different temperatures for thermal runs. Reaction time: 60 s. Feed: 0.08 g of n-octane. Each conversion and mass balance closure is an average value of three repeat runs. Temperature Conversion of nC₈ Mass Balance (° C.) (%) Closure (%) 350 0 5.47 400 0 3.12 450 0.34 7.80 Thermal Runs Using 10 wt. % Ethyl Mercaptan in n-Octane

In order to investigate EM thermal cracking, runs were developed in the CREC Catalytic Simulator with a feed containing both EM and nC₈ (10 wt. % EM/nC₈). Results are summarized in

Table 3

From these runs, it was found that there was no evidence of conversions of either nC₈ or EM at 350° C. and 400° C. and 60 s. When the temperature was increased to 450° C., there was an indication of some conversion of nC₈ but with nC₈ conversion remaining below 1%; this was consistent with the very low conversion observed for thermal cracking when pure nC₈ was fed to the reactor unit. For EM, the thermal conversion at these conditions was, however, 7.2%. This EM conversion was judged to be modest, especially considering that these experiments were developed under severe conditions (450° C., 60 s, 10 wt. % EM concentration).

The foregoing indicates that any chemical species change, as observed during the catalytic runs, is essentially the result of the ZSM-5 catalytic activity and that there is minimal influence of thermal effects. TABLE 3 Conversion of nC₈ and EM at different temperatures for thermal runs. Reaction time: 60 s. Feed: 0.08 g of 10 wt. % EM/nC₈. Each conversion and mass balance closure is an average value of three repeat runs. Standard Standard Conver- Deviation Conver- Deviation Mass Temperature sion of for nC₈ sion of for EM Balance (° C.) nC₈ (%) Conv. EM (%) Conv. Closure (%) 350 0 0 0 0 9.12 400 0 0 0 0 10.51 450 0.60 0.11 7.23 0.90 8.20

In spite of this very low thermal reaction contribution it was interesting to review the products formed, which are presented in Table 4. For a run at 450° C., 60 s, and 10 wt. % EM/nC₈ mixture, ethane, propene, and trans-butene are the detected product species with ethene being the more abundant species followed by propene and trans-butene.

It was also observed that ethene was now at higher levels than in the case of the thermal conversion with nC₈ alone: 0.61 wt. % versus 0.21 wt. %, or three times higher. This suggests that a good fraction of ethene formed in this case was originated via EM conversion, a result of intra-molecular H₂S removal. Propene, on the other hand, increased from 0.13 wt. % to 0.25 wt. %, or twice, and this points towards a mild sharing of the EM conversion via a reaction involving possibly intra-molecular H₂S removal and an alkylation step. Finally, these experiments also show the formation of trans-butene species, not observed for the nC₈ thermal conversion, and this strongly also suggests an EM conversion via inter-molecular removal of H₂S. Thus, the observed reactor products strongly support both intra-molecular and inter-molecular reactions contributing to EM conversion under thermal conditions as described with equations (1)-(3). The surprising level of olefin formation advantageously results in increased octane number in the gasoline, rather than the decrease in octane number that is experienced in conventional HDS processes.

Another surprising finding from these experiments is the absence of diethyl sulfide species, which were expected as a result of the inter-molecular dehydrosulfidation reaction step. However, it appears that this species is very reactive, being formed and consumed very quickly and remaining under non-detectable levels. This result is in marked contrast to the prior art which produces high levels of Die-S. TABLE 4 Product composition in mass fraction (wi) for thermal run. Reaction conditions: T = 450° C., t = 60 s. Feed: 0.08 g of 10 wt. % EM/nC8. Species Mass fraction (w_(i)) ethene 0.0061 propene 0.0025 trans-butene 0.0017 ethyl mercaptan 0.0928 n-octane 0.8947 hydrogen sulfide 0.0040

While H₂S could not be detected with the FID, it was positively identified using the GC-MSD and quantified via element balances.

Catalytic Runs: Conversions of Ethyl Mercaptan

FIG. 1 reports the EM conversions for the concentration of 10 wt. % of EM in nC₈. It can be observed that the EM conversion increases progressively with reaction time in the 10-60 s range. In addition there is an increase of catalytic dehydrosulfidation with temperature.

Furthermore, and in order to explore the reaction order of the EM reaction, the EM in nC₈ was decreased to 5 wt. % keeping the C/0 at 2.5. Results of these experiments are reported in FIG. 2. The lack of dependence of the EM conversions, at various temperatures and reaction times, with the EM concentration (comparison between FIGS. 1 and 2) suggests an overall first order reaction for the catalytic EM conversion.

Catalytic Runs: Conversions of n-Octane

Conversions of n-octane are reported in FIGS. 3, 4 and 5. These figures contain results of experiments developed using 100 wt. % of nC₈, 95 wt. % of nC₈ (5 wt. % of EM) and 90 wt. % of nC₈ (10 wt. % of EM). It can be noticed that the runs showed an expected and progressive increase of the nC₈ conversion with reaction time and temperature. Conversions levels, however, for the nC₈ remained at much lower levels than the EM conversion. For instance, at 450° C., 60 s and C/O=2.5 the observed nC₈ conversion was 22.87% (0.5 S.D.) versus 50.85% (0.91 S.D.) for EM at the same conditions.

These results show that there is a strong competition for the acid sites of the ZSM-5 catalyst promoting both dehydrosulfidation and catalytic cracking. It appears that, given the significant differences in gas phase concentrations between nC₈ and EM (between 10 and 20 times), there is either a greater affinity of EM adsorption versus nC₈ adsorption or, alternatively, a much faster intrinsic rate of EM dehydrosulfidation versus the one for nC₈ cracking.

Catalytic Runs: Product Distribution

In the catalytic runs using pure nC₈, the obtained amount of propene was always higher than the amount of ethene, with this being a typical characteristic of catalytic cracking. Trans-butene was only formed at 400° C. and 450° C. and 40 and 60 s and the amounts produced were very low. TABLE 5 Product distribution of key species for reactions using 10 wt. % EM/nC8. Compositions expressed in wt. % Time (min): 10 20 40 60 350° C. ethene 0.07 0.21 0.46 0.80 propene 0.41 0.71 1.18 1.47 trans-butene 0.00 0.04 0.10 0.19 H₂S 0.34 0.58 1.10 1.42 400° C. ethene 0.20 0.37 0.66 0.80 propene 0.78 1.46 2.80 3.83 trans-butene 0.04 0.08 0.16 0.22 H₂S 0.63 1.12 1.99 2.58 450° C. ethene 0.64 1.06 1.55 1.99 propene 1.35 2.71 4.88 7.14 trans-butene 0.14 0.25 0.40 0.51 H₂S 1.06 1.92 2.86 3.78

Product composition, as reported in Table 5, for the catalytic runs using mixture of EM and nC₈ shows trans-butene levels produced being higher than those obtained from runs using pure nC₈. This indicates that there is formation of trans-butene via the inter-molecular de-hydrosulfidation reaction. Hydrogen sulfide amounts increase with residence time and temperature, with this being in agreement with the rise of EM conversion with these two parameters. Furthermore, the amounts of ethene observed were also higher for the catalytic runs using mixtures of EM and nC₈ than for the corresponding cases using pure nC₈. This gives a good indication of formation of ethene via the intra-molecular dehydrosulfidation reaction.

The ratios of tC₄ ⁼/C₂ ⁼ (trans-butene/ethene), which provide valuable insights on the relative importance of inter-molecular and intra-molecular catalytic dehydrosulfidation, are reported in Table 6. TABLE 6 Trans-butene to ethene (tC₄ ⁼/C₂ ⁼) ratios for catalytic experiments. T = 350° C. nC₈ EM (5 wt. %) EM (10 wt. %) Time (min) tC₄ ⁼/C₂ ⁼ tC₄ ⁼/C₂ ⁼ tC₄ ⁼/C₂ ⁼ 10 — 0 0 20 — 0.1976 0.1839 40 — 0.3087 0.2187 60 — 0.322 0.2373 T = 400° C. nC8 EM (5 wt. %) EM (10 wt. %) Time (min) tC₄ ⁼/C₂ ⁼ tC₄ ⁼/C₂ ⁼ tC₄ ⁼/C₂ ⁼ 10 — 0.3322 0.2031 20 0 0.395 0.2074 40 0.117 0.3934 0.244 60 0.2357 0.4197 0.274 T = 450° C. nC8 EM (5 wt. %) EM (10 wt. %) Time (min) tC₄ ⁼/C₂ ⁼ tC₄ ⁼/C₂ ⁼ tC₄ ⁼/C₂ ⁼ 10 0 0.324 0.216 20 0.1059 0.372 0.2354 40 0.1967 0.394 0.2544 60 0.2335 0.4128 0.2559

It can be noticed that catalytic nC₈ conversion leads to tC₄ ⁼/C₂ ⁼ ratios much smaller than the ones observed for the catalytic conversion of EM-nC₈ mixtures. For instance for the catalytic conversion of nC₈ at 350° C., this ratio is zero with no tC₄ ⁼ being detected; this ratio also remains at small levels at 400° C. and 450° C. Thus, it can be argued on this basis that at 350° C. olefin dimerization under these conditions is negligible, becoming somewhat more prevalent at 400° C. and 450° C.

On the other hand, the tC₄ ⁼/C₂ ⁼ ratios increase considerably when the EM/nC₈ mixtures are contacted with the catalyst and this difference suggests an increased influence of the competitive conversion of EM via the inter-molecular path, with more tC₄ ⁼ species being formed. There is evidence that increasing the temperature leads to a more significant inter-molecular EM conversion. The surprising levels of olefin formation observed in the presence of the catalyst is another advantage of this process. TABLE 7 Product composition for catalytic dehydrodesulfidation in the CREC Catalytic Simualtor. T = 450° C., t = 60 s, C/O = 2.5, Reactants: 5% of EM in nC₈ Chemical Species Compositon (wt %) Ethene 0.015 Propene 0.0817 Iso-butane 0.028 n-butane 0.053 t-butene 0.00626 2-methybutane 0.0105 1-pentene 0.0014 n-pentane 0.0137 Ethyl-mercaptan 0.0266 Cis-1,2-dimethycyclopropane 0.052 2-methy-pentane 0.0026 1-hexene 0.0014 n-hexane 0.0023 Benzene 0.0015 Toluene 0.0093 n-octane 0.794 Xylenes 0.0088 Hydrogen sulfide 0.0232

Referring to Table 7, the only reaction contributing significantly to the formation of C₂ and C₄ species is the catalytic dehydrosulfidation of gasoline with minimum contribution of gasoline cracking (less than 5%). The C₂ and C₄ species produced are predominantly olefins (eg: ethene and propene) that have significant petrochemical and commercial value. With nC₈ as a feed stock, there were essentially no C₆ aromatics (eg: benzene, toluene, xylene) produced, and no coke formation observed. It is also worthwhile to note that essentially no diethyl sulfide was produced, contrary to the prior art and providing evidence of the desired high selectivity towards hydrogen sulfide.

In a commercial process where the hydrocarbon stream is a mixture of parafins, olefins and aromatics, there is some coke formation expected. In a preferred embodiment, the process is developed concurrently with catalyst regeneration where the coke, deactivating the catalyst, is combusted in a dense phase fluidized bed and the overall process is envisioned with the catalyst continuously circulating between two twin reactors: a solids transport reactor for catalytic dehydrosulfidation of gasoline, preferably a downer or riser reactor; and, a dense phase fluidized bed for coke combustion and catalyst regeneration. A schematic representation of an embodiment of this process is provided in FIG. 7 with a downer reactor. FIG. 7 will be described in greater detail hereinafter.

Furthermore, the process of catalytic dehydrosulfidation should not be considered limited to mercaptan dehydrosulfidation. This is quite critical given that there is a significant fraction of sulfur contained in gasoline in the form of thio-aromatic compounds. For instance, one could envision intermolecular dehydrosulfidation of one mercaptan molecule reacting with one thio-aromatic molecule such as benzo-thiophene. In this case the products of dehydrosulfidation will yield molecules with a critical molecular diameter larger than the narrow 5.4 Å channels of the ZSM5 zeolite, with these coke precursors remaining trapped in the ZSM-5 zeolite porous network and increasing the need for catalyst regeneration.

Given the foregoing experimental results and analysis, the following general observations and conclusions can be reached:

-   1. Gas phase and catalytic phase residence times should be very     close to achieve the desired high conversion of organo-sulfur     compounds and selectivity to H₂S; -   2. The catalytic reactor operates under close to isothermal     conditions; -   3. The EM reaction exhibits first-order kinetics; and, -   4. The series-parallel reaction network dominates the progress of     the catalytic dehydrosulfidation reaction. This implies long gas     phase residence times (>7 s) to achieve the desired conversion     levels.     Reactor Design Calculations

The CREC Catalytic Simulator design equation can, considering the various special design features of this unit (batch operation, well fluidized catalyst contained between two grids, isothermal operation, high re-circulation of chemical species) be expressed as follows: $\begin{matrix} {\frac{\mathbb{d}C_{EM}}{\mathbb{d}t} = {r_{EM}\frac{W_{c}}{V_{r}}}} & (4) \end{matrix}$ where W_(c) is the weight of the zeolite crystallites and V_(r) is the reactor volume (cm³).

If a first order overall kinetics for EM conversion, as suggested by the experimental data in the CREC Catalytic Simulator, is assumed it results: $\begin{matrix} {\frac{\mathbb{d}C_{EM}}{\mathbb{d}t} = {{kC}_{EM}\frac{W_{c}}{V_{r}}}} & (5) \end{matrix}$

Equation 5 can be written in terms of ethyl-mercaptan fractional conversion and following integration it gives: $\begin{matrix} {{- {\ln\left( {1 - X_{EM}} \right)}} = {\frac{k\quad W_{c}}{V_{r}}t}} & (6) \end{matrix}$ where X_(EM) is the fractional conversion of EM.

A linear regression can be applied to equation 6 to find the slope (kW_(c)/V_(r)). Knowing that the loading of zeolite crystallites was 0.2 grams and that the CREC Catalytic Simulator volume was 52 cm³, the value of k may be calculated.

Table 8 reports the obtained values for k with 95% confidence bounds. TABLE 8 Values of the kinetic constant at 350° C., 400° C. and 450° C. Temperature k (cm³/g_(cryst) · s) ±C.L. (95%) R² 350° C. 1.11 0.0424 0.9624 400° C. 2.05 0.1169 0.8574 450° C. 3.88 0.1956 0.9184

An Arrhenius relationship k=k₀exp(−E/RT) can be adopted to express the temperature effect. This equation can be modified using re-parametrization to reduce cross-correlation between the activation energy and the pre-exponential factor. This is accomplished employing a central temperature (T₀), the average value of the three temperatures used in this research. Thus, equation 6 becomes, $\begin{matrix} {X_{{EM},{model}} = {1 - {\exp\left\lbrack {{- \frac{W_{c}}{V_{r}}}{t\left( {k_{o}^{\prime}\quad{\exp\left( {{- \frac{E}{R}}\left( {\frac{1}{T} - \frac{1}{T_{o}}} \right)} \right)}} \right)}} \right\rbrack}}} & (7) \end{matrix}$

A non-linear regression with minimization of residuals Σ(X_(EM,exp)−X_(ETM,model))² was developed using a MatLab program and initial guesses for k as reported in Table 1.

The resulting kinetic parameters with their confidence limits at the 95% were activation energy (E=46.66±34.84 KJ/mole) and pre-exponential factor (k₀ ¹=2.136 cm³/g_(cryst)·s).

The following design equations apply for first order kinetics in the CREC Catalytic Simulator EMConversion=1−exp(−kδ′ _(cat) t)  (8) and k=k ₀ ¹ exp[−E/R(1/T−1/T ₀)]  (9)

with k₀ ¹ representing the kinetic constant at 400° C., E the energy of activation (KJ/mole), t the reaction time (s) and δ′_(cat) the catalyst reactor density (g_(cryst)/cm³ of reactor).

Using equation (9), with T=450° C. and T₀=400° C. results in an observed kinetic constant k for the CREC Catalytic Simulator of 3.809 cm³/g_(cryst)·s. With 60 seconds total reaction time, C/O=2.5 and T=450° C. a conversion of 56% is calculated for the CREC Catalytic Simulator using the above parameters. This compares favourably with the actual conversion observed in FIG. 1.

The design of a full-scale reactor for use in the process of the present invention can be based on the following assumptions:

-   1. An essentially plug flow pattern for both gas and solid phases; -   2. A negligible velocity difference between gas and particles, thus     gas and particles can be assumed flowing through the unit with the     same velocities; -   3. A negligible variation of particle and solid velocities across     the reactor cross section; -   4. An isothermal operation, given the relatively low total enthalpy     changes involved in the dehydrosulfidation process and the high     radial mixing of solid particles; and, -   5. Similar residence times for catalyst and reactants in the range     of less than 30 seconds.

Given the above constraints, there are two basic types of solids transport reactors that can be chosen: risers and downers. Both types of reactors circulate the catalyst continuously, but the flow direction in risers is upward, whereas the flow direction in downers is downward.

On this basis the EM balance for the reactor can be expressed under steady state operation as follows: $\begin{matrix} {{u_{g}\frac{\mathbb{d}C_{EM}}{\mathbb{d}z}} = {r_{EM}\rho_{c}}} & (10) \end{matrix}$ where ρ_(c) is the catalyst reactor density (expressed as mass of zeolite crystallites per unit reactor volume) and z is the axial reactor length in m and u_(g) represents the gas interstitial velocity in the reactor in m/s.

If a first order overall kinetics for EM conversion, as suggested by the experimental data in the CREC Catalytic Simulator, is assumed it results: $\begin{matrix} {\frac{\mathbb{d}C_{EM}}{\mathbb{d}t_{R}} = {{- k}\quad C_{EM}\rho_{c}}} & (11) \end{matrix}$ where t_(R) represents the gas phase residence time.

Equation 11 can be alternatively written in terms of ethyl-mercaptan fractional conversion and following integration it gives: −ln(1−X _(EM))=kρ _(c)t_(R)  (12)

Thus, one can postulate that the EM conversion in the reactor for catalytic dehydrosulfidation of gasoline can be expressed as $\begin{matrix} {X_{{EM},{model}} = {1 - {\exp\left\lbrack {{- \rho_{c}}{t_{R}\left( {k_{o}^{\prime}{\exp\left( {{- \frac{E}{R}}\left( {\frac{1}{T} - \frac{1}{T_{o}}} \right)} \right)}} \right)}} \right\rbrack}}} & (13) \end{matrix}$

By using in equation 12 an EM conversion of 95% with k=3.809 cm³/g_(cryst)·s (for the parameters T=450° C., E=46.66±34.84 KJ/mole, k₀ ¹=2.13±0.05 cm³/g_(cryst)·s), it can be determined that ρ_(c)t_(R) is a constant equal to 0.786 g_(cryst)·s/cm³. Thus, one could secure the desired 95% EM conversion by selecting operating conditions that satisfy the relationship ρ_(c)t_(R)=0.786 g_(cryst)·s/cm³.

For example, if one chooses a residence time in the reactor of 12 seconds (gas superficial velocity=1.8 m/s, total reactor length=21.6 m) the required catalyst reactor density is 0.063 g_(cryst)/cm³ reactor. Given than one should expect that, in a catalyst particle, the ratio between the total weight (matrix+crystallites) of 60 micron pellets and the weight of zeolite crystallites to be about 3 g_(particle)/g_(cryst), the overall catalyst density is about 0.189 g_(particle)/cm³ reactor. Even more, assuming that the density of the pellet is about 1 g_(particle)/cm³ catalyst, this yields an apparent volumetric concentration of 0.189 cm³ catalyst/cm³ reactor. Finally, and considering that an estimated 10% of the gas volumetric follow in the downer is inert gas used to carry the particles from the regenerator via pneumatic transport, a 0.9 correction is needed. This results in a volumetric particle concentration of 0.17 cm³ particle/cm³ reactor or 17% for achieving 95% EM conversion.

This high volumetric particle concentration is not achievable in a riser reactor, given the constraints imposed on the volumetric particle concentration by the choking of the suspension. Referring to Kerry and Knowlton (“Wall Solid Upflow and Down Flow Regimes in Risers for Group A Solids”, Proceedings of the Circulating Fluidized Bed Technology VII Conference, Niagara Falls, 2002, eds J. Grace., J. Zhu and H. de Lasa, pp. 310-316), FIG. 6 shows a choking velocity correlation for 67 micron FCC particles of group A; these particles are very similar to the ones considered for use in the present invention. At a superficial gas velocity of 2.8 m/s, the highest volumetric particle concentration before onset of choking is 3.5%. At a superficial gas velocity of 1.8 m/s, the highest volumetric particle concentration before onset of choking is 1.1%. This is nearly an order of magnitude lower than the volumetric particle concentration of 17% needed to achieve 95% EM conversion. A riser reactor is therefore unsuitable for achieving the desired combination of process conditions.

However, this high volumetric particle concentration is perfectly achievable in a downer reactor as reported by Jin, et al. (Y. Jin, Yu Zheng, F. Wei “State of the Art Review of Downer Reactors”, Proceedings of the Circulating Fluidized Bed Technology VII Conference, Niagara Falls, 2002, eds J. Grace, J. Zhu and H. de Lasa, pp. 40-60). Particles move in a downward direction in downers, the same direction as gravity, not against gravity as in a riser unit. In a downer unit there is no possible choking of the suspension, so very high solids/gas loading ratios and volumetric particle concentrations may be achieved. A practical upper limit of volumetric particle concentration in a downer is 40%, which permits the 17% value calculated above for 95% EM conversion.

The relationship between catalyst reactor density and residence time is illustrated in FIG. 6 for ρ_(c)t_(R)=0.786 g_(cryst)·s/cm³, which corresponds to 95% conversion at T=450° C. with k=3.809 cm³/g_(cryst)·s. Limit 1 represents the practical limitation on residence time that results in a practical reactor length. A residence time limit of 20 seconds corresponds to a reactor length of 36 m. A preferred reactor length for this process is between about 14 and 22 metres. Limit 2 represents the maximum catalyst reactor density that can be practically utilized in a downer reactor while still satisfying the fluid dynamics assumptions listed above. Limit 2 corresponds to about 0.12 g_(cryst)/cm³ reactor, or about 40% volumetric particle concentration in the reactor. For comparison, a volumetric particle concentration limit of 3.5% in riser reactors before the onset of choking corresponds to 0.001 g_(cryst)/cm³, which places Limit 2 far to the right of Limit 1. Therefore, 95% conversion is simply not achievable in a riser reactor. Thus, the downer reactor is the only practical reactor system that provides the unique range of conditions (high solid concentration, moderate gas superficial velocity compatible with a proper total reactor length) desired to achieve the targets for conversion and selectivity.

Calculations were conducted using the above equations and are summarized in Table 9, which discloses design parameters for several embodiments of the process according to the present invention. As can be seen from the Table, high conversions can be achieved with reasonable residence times and reactor lengths. The volumetric concentration of catalyst in all cases is suitable for a downer reactor. By using the parameters disclosed in Table 9, high selectivity towards hydrogen sulfide should also be achieved in each embodiment. TABLE 9 Catalytic Dehydrosulfidation Conditions in a Downer Reactor at various conversions and gas residence times. T = 450° C., k = 3.809 cm³/g_(cryst) · s. Volumetric Particle Reactor Conversion t_(R) Concentration ρ_(c) ρ_(c) t_(R) Length (%) (s) (%) g_(cryst)/cm³ g_(cryst) · s/cm³ (m) 70 9 9.45 0.035 0.315 16.2 80 9 12.6 0.0468 0.412 16.2 90 9 18.1 0.067 0.603 16.2 95 9 23.4 0.0868 0.781 16.2 70 12 7.1 0.0263 0.3156 21.6 80 12 9.5 0.0352 0.422 21.6 90 12 13.6 0.0505 0.606 21.6 95 12 17.6 0.0655 0.786 21.6 70 15 5.69 0.0211 0.316 27 80 15 7.58 0.0281 0.421 27 90 15 10.91 0.0404 0.606 27 95 15 14.17 0.0525 0.787 27

Referring to FIG. 7, a process of the present invention is schematically represented. Streams of untreated gasoline 5 and ZSM-5 catalyst 10 are introduced into a downer reactor 15 near the top where the streams are co-mingled. As the co-mingled streams pass through the downer reactor, sulfur of organo-sulfur compounds in the untreated gasoline is catalytically converted to hydrogen sulfide.

The co-mingled stream exits the downer reactor 15 at the bottom and is fed to a first cyclone separator 25 where treated gasoline is separated from catalyst. Treated gasoline is recovered in treated gasoline stream 30 whereas separated catalyst stream 35 is combined with a first carrier fluid stream 40 (e.g. an inert gas or steam) and fed to a dense phase fluidized catalyst regenerator 45.

An air stream 50 is introduced into the regenerator 45 and any coke is combusted in the presence of the air to regenerate the catalyst. Combustion gases are removed from the top of the regenerator in combustion gas stream 60. A regenerated catalyst stream 65 exits the regenerator at the bottom and is combined with a second carrier fluid stream 70 (e.g. an inert gas or steam) and fed to a second cyclone separator 75 where most (about 90% or more) of the carrier fluid is separated from the catalyst. Separated carrier fluid is recovered in separated carrier fluid stream 80 and catalyst is reintroduced into the downer reactor in catalyst stream 10.

Other advantages which are inherent to the structure are obvious to one skilled in the art. The embodiments are described herein illustratively and are not meant to limit the scope of the invention as claimed. Variations of the foregoing embodiments will be evident to a person of ordinary skill and are intended by the inventor to be encompassed by the following claims. 

1. A process for removing sulfur from a hydrocarbon stream comprising an organo-sulfur compound, the process comprising: a) providing a downer reactor; b) providing a hydrocarbon stream comprising an organo-sulfur compound in the downer reactor at a temperature in the range of 300 to 500° C.; c) providing a catalyst comprising a zeolite that promotes removal of sulfur from the organo-sulfur compound in the downer reactor; d) contacting the hydrocarbon stream with the catalyst; and, e) converting at least 70% of the organo-sulfur compound to hydrogen sulfide in the downer reactor.
 2. A process according to claim 1, wherein the catalyst has a residence time in the downer reactor and wherein the residence time is from 7 second to 30 seconds.
 3. A process according to claim 1, wherein at least 75% of the organo-sulfur compound is converted to hydrogen sulfide in the downer reactor.
 4. A process according to claim 1, wherein the temperature is from about 375° C. to about 500° C.
 5. A process according to claim 1, wherein the catalyst is present in the downer reactor in a volumetric particle concentration and wherein the volumetric particle concentration is from about 5 to about 40%.
 6. A process according to claim 1, wherein the hydrocarbon stream has a fluid velocity in the downer reactor and wherein the fluid velocity is between 0.5 and 4.0 m/s.
 7. A process according to claim 1, wherein the downer reactor has a length of between 5 and 40 meters.
 8. A process according to claim 1, wherein the downer is operated such that the product of catalyst reactor density and residence time is between 0.30 and 1.44 g_(cryst)·s·cm⁻³.
 9. A process according to claim 1, wherein the catalyst is a microporous hydrophobic zeolite having acidic sites and a tortuous pore path.
 10. A process according to claim 1, wherein the catalyst comprises ZSM-5.
 11. A process according to claim 1, wherein the organo-sulfur compound is present in the hydrocarbon stream at a concentration of between 10 and 10000 ppm.
 12. A process according to claim 1, wherein the organo-sulfur compound comprises one or more sulfur atoms and from 1 to 12 carbon atoms.
 13. A process according to claim 1, wherein the organo-sulfur compound comprises a mercaptan, a thio-ether, a thio-ketone, a thio-aromatic, a thio-paraffin, or a combination thereof.
 14. A process according to claim 1, wherein the organo-sulfur compound comprises methyl mercaptan, ethyl mercaptan, or diethyl sulfide.
 15. A process according to claim 1, wherein the hydrocarbon stream comprises a hydrocarbon having from 1 to 20 carbon atoms.
 16. A process according to claim 1, wherein the hydrocarbon stream comprises n-octane.
 17. A process according to claim 1, wherein the hydrocarbon stream is gasoline.
 18. A process according to claim 1, wherein the process further comprises circulating the catalyst between the downer reactor and a regenerator.
 19. A process for removing sulfur from a hydrocarbon stream comprising an organo-sulfur compound, the process comprising: a) providing a solids transport reactor; b) providing a hydrocarbon stream comprising an organo-sulfur compound in the solids transport reactor at a temperature in the range of 300 to 500° C.; c) providing a catalyst comprising a zeolite that promotes removal of sulfur from the organo-sulfur compound in the solids transport reactor; the catalyst having a catalyst reactor density from 0.02 to 0.12 g_(cryst)·cm⁻³; d) contacting the hydrocarbon stream with the catalyst; and, e) converting at least 70% of the organo-sulfur compound to hydrogen sulfide in the solids transport reactor in a residence time from 7 to 20 seconds, wherein the product of catalyst reactor density and residence time is from 0.30 to 1.44 g_(cryst)·s·cm⁻³.
 20. A process for removing sulfur from a hydrocarbon stream comprising an organo-sulfur compound, the process comprising: a) providing a downer reactor having a length of from 10 to 36 meters; b) providing a hydrocarbon stream comprising gasoline and comprising 10 to 10000 ppm of an organo-sulfur compound comprising a mercaptan in the downer reactor at a temperature in the range of 300 to 500° C.; c) providing a catalyst comprising zeolite ZSM-5 in the downer reactor at a volumetric particle concentration of from 5% to 40%; d) contacting the hydrocarbon stream with the catalyst; and, e) converting at least 90% of the organo-sulfur compound to hydrogen sulfide in the downer reactor in a residence time of between 7 and 20 seconds. 